Process for the preparation of a lubricant

ABSTRACT

Process to prepare a lubricant having a dynamic viscosity at −35° C. of below 5000 cP by performing the following steps: a) contacting a feed containing more than 50 wt % wax in the presence of hydrogen with a catalyst comprising a Group VIII metal component supported on a refractory oxide carrier, and b) contacting the effluent of step (a) with a catalyst composition comprising a noble Group VIII metal, a binder and zeolite crystallites of the MTW type to obtain a product having a lower pour point than the effluent of step (b) and having a viscosity index greater than 120, and (c) adding a pour point depressant additive to the base oil as obtained in step (b).

PRIORITY CLAIM

The present application claims priority to European Patent ApplicationNo. 02293035.8 filed 9 Dec. 2002.

The invention is directed to a process to prepare a lubricant having adynamic viscosity at −35° C. of below 5000 cP by the addition of a pourpoint depressant additive to a base oil obtained by means of a processinvolving catalytic dewaxing.

In GB-A-1429494 a process for the preparation of a lubricating oil isdisclosed wherein the heavy fraction of a hydrocracked slack wax issolvent dewaxed, using a mixture of solvents. The resulting base oil hasa very high viscosity index of up to about 155.

The disadvantage of the process is that the response to pour pointdepressant additives is not sufficient to prepare the desired lubricant.

EP-A-324528 describes a process for the preparation of a lubricatingbase oil wherein a slack wax is first hydrocracked. The effluent of thehydrocracker is subsequently catalytically dewaxed in the presence of acatalyst containing a zeolite crystallite of the MFI type. The resultantbase oil had a viscosity index (VI) of 134 and a pour point of −44° C.The yield of base oil related to the starting slack wax feed is howeverlow.

It is an object of the present invention to provide a process for thepreparation of lubricants wherein the prepared base oils have aviscosity index greater than 120 and a favorable response to pour pointdepressant additives. It is furthermore an object of the presentinvention to provide a process wherein the yield to base oils isimproved starting from a slack wax feed.

This object is achieved with the following process. Process to prepare alubricant having a dynamic viscosity at −35° C. of below 5000 cP byperforming the following steps:

(a) contacting a feed containing more than 50 wt % wax in the presenceof hydrogen with a catalyst comprising a Group VIII metal componentsupported on a refractory oxide carrier, and

(b) contacting the effluent of step (a) with a catalyst compositioncomprising a noble Group VIII metal, a binder and zeolite crystallitesof the MTW type to obtain a product having a lower pour point than theeffluent of step (b) and having a viscosity index greater than 120, and(c) adding a pour point depressant additive to the base oil as obtainedin step (b).

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 presents a graph of the viscosity index as a function of achievedpour point of an oil fraction.

FIG. 2 presents a graph of the oil yield as a function of achieved pourpoint of an oil fraction.

FIG. 3 presents a graph that shows WABT as a function of pour point ofdewaxed oil.

FIG. 4 presents a graph that shows the influence of wax conversion in aprocess step on overall base oil yield.

FIG. 5 presents a graph of the dynamic viscosity at −35° C. of anintermediate product that was catalytically and solvent dewaxed atvarious pour points between −20 and −37° C.

Applicants found that when a pour point depressant additive is added tothe base oils as prepared by the process according to the presentinvention a significantly larger reduction in pour point is observed ascompared to when the same additive is added to a prior art solventdewaxed base oil. In addition also the dynamic viscosity as measured at−35° C. according to ASTM D 2983 is significantly more reduced afteradding the additive. This is very advantageous. For example high tierlubricants having a dynamic viscosity at −35° C. of below 5000 cP may beprepared with a base oil as prepared according to the present inventionhaving a kinematic viscosity at 100° C. of between 4.5 and 5.5 cSt and apour point of between −18 and −35° C. and more preferably between −18and −30° C. and most preferably between −23 and −30° C. The fact thatthe pour point of the base oil can be relatively high is veryadvantageous in terms of base oil yield and the viscosity index of thebase oil. The viscosity index of the dewaxed base oil having the abovepour point range may be advantageously between 135 and 150.

The feed to step (a) may contain between 0-3000 ppm sulphur and between0-150 ppm nitrogen. The feed may be a synthetic wax, such as for examplederived from a Fischer-Tropsch process, or even a slack wax.

Slack wax can be obtained from either a hydrocracked oil or a solventrefined oil. Hydrocracking is preferred because that process can alsoreduce the nitrogen content to low values. With slack wax derived fromsolvent refined oils, de-oiling can be used to reduce the nitrogencontent. The oil content of the slack wax feed may be between 0 and 50wt %. Slack wax having a high oil content may be obtained as the directby-product of a solvent dewaxing process. Examples of suitable solventdewaxing processes are described in Lubricant Base Oil and WaxProcessing, Avilino Sequeira, Jr, Marcel Dekker Inc., New York, 1994,Chapter 7.

The content of aromatic compounds in the waxy feed will suitably bebetween 3 and 40 wt %. The slack wax feed preferably has a very highviscosity index, more preferably in the range of from 140 to 200. Theviscosity index of the feed will determine in part the viscosity indexof the resulting lubricating base oil, which are preferably between 120and 180. The VI will depend on the oil content and the starting materialfrom which the wax has been prepared. Optionally, hydrotreating of theslack wax feed, prior to performing step (a) can be carried out to lowerthe nitrogen content thereof.

The feed used in the process according to this invention contains morethan 50 wt % wax, preferably between 50 and 100 wt % wax and morepreferably between 70 and 100 wt % wax.

The wax content as used in the description is measured according to thefollowing procedure. 1 weight part of the to be measured oil fraction isdiluted with 4 parts of a (50/50 vol/vol) mixture of methyl ethyl ketoneand toluene, which is subsequently cooled to −27° C. in a refrigerator.The mixture is subsequently filtered at −27° C. The wax is removed fromthe filter and weighed. If reference is made to oil content a wt % valueis meant which is 100% minus the wax content in wt %.

Step (a) may be performed using well known hydrocracking processes suchas described in Lubricant Base Oil and Wax Processing, Avilino Sequeira,Jr, Marcel Dekker Inc., New York, 1994, Chapter 6 and especially pages121-131. Catalysts for use in step (a) typically comprise an acidicfunctionality and a hydrogenation/dehydrogenation functionality.Preferred acidic functionality's are refractory metal oxide carriers.Suitable carrier materials include silica, alumina, silica-alumina,zirconia, titania and mixtures thereof. Preferred carrier materials forinclusion in the catalyst for use in the process of this invention aresilica, alumina and silica-alumina.

The catalyst may comprise a Group VIII noble metal, for examplepalladium and more preferably platinum. Such a catalyst is preferredwhen the feed to step (a) is substantially free of contaminants such assulphur, i.e. less than 10 ppm or nitrogen, i.e. less than 10 ppm. Suchwax containing feeds may advantageously be obtained from the synthesisproduct of a Fischer-Tropsch reaction. The catalyst may comprise thenoble metal in an amount of from 0.005 to 5 parts by weight, preferablyfrom 0.02 to 2 parts by weight, per 100 parts by weight of carriermaterial. A particularly preferred catalyst comprises platinum in anamount in the range of from 0.05 to 2 parts by weight, more preferablyfrom 0.1 to 1 parts by weight, per 100 parts by weight of carriermaterial. Examples of possible noble metal catalysts are C-624 ofCriterion Catalyst Company or those described in for exampleEP-A-666894.

The process according to the invention is particularly suited to processwaxy feedstocks which have a mineral oil source, such as for exampleslack wax. These mineral oil derived feeds will comprise substantialamounts of nitrogen- and sulphur-containing compounds. When such a feedis used the catalyst in step (a) preferably comprises a Group VIB metaland a non-noble Group VIII metal. Possible combinations of one or moreof the metals cobalt, iron and nickel, and one or more of the metalschromium, molybdenum and tungsten are preferred. Especially preferredcatalysts for use in treating such feeds comprise, in combination,cobalt and molybdenum, nickel and tungsten and nickel and molybdenum.Typically, the catalyst comprises from 10 to 100 parts by weight of theGroup VIB metal, if present, preferably from 25 to 80 parts weight, per100 parts by weight of carrier. The Group VIII non-noble metal istypically present in an amount of from 3 to 100 parts by weight, morepreferably from 25 to 80 parts by weight, per 100 parts by weight ofcarrier. If desired, applying a halogen moiety, in particular fluorine,or a phosphorous moiety to the carrier, may enhance the acidity of thecatalyst carrier. Examples of suitable hydrocracking/hydroisomerisationprocesses and suitable catalysts are described in GB-A-1493620,WO-A-9941337, U.S. Pat. No. 5,370,788, EP-A-537969, U.S. Pat. No.5,292,989, WO-A-0014179, EP-A-532118, EP-A-666894 and EP-A-776959.

In case of a mineral oil feed to step (a) preferably use is made of asulphided catalyst based on a Group VIB metal and a non-noble Group VIIImetal as described above. Sulphidation of the catalyst may be effectedby any of the techniques known in the art, such as ex-situ or in-situsulphidation. For example, sulphidation may be effected by contactingthe catalyst with a sulphur-containing gas, such as a mixture ofhydrogen and hydrogen sulphide, a mixture of hydrogen and carbondisulphide or a mixture of hydrogen and a mercaptan, such asbutylmercaptan. Alternatively, sulphidation may be carried out bycontacting the catalyst with hydrogen and sulphur-containing hydrocarbonoil, such as sulphur-containing kerosene or gas oil. The sulphur mayalso be introduced into the hydrocarbon oil by the addition of asuitable sulphur-containing compound, for example dimethyldisulphide ortertiononylpolysulphide.

The feedstock will preferably comprise a minimum amount of sulphur inorder to keep the catalyst in a sulphided state. Preferably at least 200ppm sulphur and more preferably at least 700 ppm sulphur is present inthe feed. It may therefore be necessary to add additional sulphur, forexample as dimethylsulphide, or a sulphur containing co-feed to the feedof step (a) if the feed contains a lower level of sulphur. Examples ofmineral oil feeds, which contain lower levels of sulphur, are slackwaxes obtained from oil, which has been obtained in a hydrocrackingprocess. Such slack waxes may contain between 10-200 ppm sulphur.

A preferred catalyst for use in step (a) is a sulphidedhydrodesulphurisation catalyst comprising nickel and tungsten on anacidic amorphous silica-alumina carrier. Such a catalyst is preferablyfree of any halogen. The sulphided hydrodesulphurisation catalyst has arelatively high hydrodesulphurisation activity. With relatively highactivity is here meant a considerably higher activity when compared tostate of the art nickel/tungsten containing catalysts based on asilica-alumina carrier. Preferably the hydrodesulphurisation activity ofthe catalyst is higher than 30% and more preferably below 40%, and mostpreferably below 35%, wherein the hydrodesulphurisation activity isexpressed as the yield in weight percentage of C₄-hydrocarbon crackingproducts when thiophene is contacted with the catalyst under standardhydrodesulphurisation conditions. The standard conditions consists ofcontacting a hydrogen/thiophene mixture with 200 mg of a 30-80 meshsulphided catalyst at 1 bar and 350° C., wherein the hydrogen rate is 54ml/min and the thiophene concentration is 6 vol % in the total gas feed.

Catalyst particles are to be used in the test are first crushed andsieved through a 30-80 mesh sieve. The catalyst is then dried for atleast 30 minutes at 300° C. before loading 200 mg of dried catalyst intoa glass reactor. Then the catalyst is pre-sulphided by contacting thecatalyst for about 2 hours with an H₂S/H₂ mixture, wherein the H₂S rateis 8.6 ml/min and the H₂ rate is 54 ml/min. The temperature during thepre-sulphiding procedure is raised from room temperature, 20° C., to270° C. at 10° C./min and held for 30 minutes at 270° C. before raisingit to 350° C. at a rate of 10° C./min.

During pre-sulphiding nickel and tungsten oxides are converted to theactive metal sulphides. After pre-sulphiding the H₂S flow is stopped andH₂ is bubbled at a rate of 54 ml/min through two thermostatted glassvessels containing thiophene. The temperature of the first glass vesselis kept at 25° C. and the temperature of the second glass vessel is keptat 16° C. As the vapour pressure of thiophene at 16° C. is 55 mmHg, thehydrogen gas that enters the glass reactor is saturated with 6 vol %thiophene. The test is performed at 1 bar and at a temperature of 350°C. The gaseous products are analysed by an online gas liquidchromatograph with a flame ionisation detector every 30 minutes for fourhours.

In order to obtain a reproducible value for the hydrodesulphurisationactivity the test values as obtained by the above method are correctedsuch to correspond to the hydrodesulphurisation activity of a referencecatalyst. The reference catalyst is the commercial C-454 catalyst asobtainable at the date of filing of Criterion Catalyst Company (Houston)and its reference hydrodesulphurisation activity is 22 wt % according tothe above test. By testing both the reference catalyst (“test C-454”)and the test catalyst (“measured val”) one can easily calculate aconsistent actual hydrodesulphurisation activity according to the abovetest with the below equation:Actual activity=“measured val”+((22−“test C-454”)/22)*“measured val”

The hydrodesulphurisation activity of the nickel/tungsten catalyst canbe improved by using chelating agents in the impregnation stage of thepreparation of the catalyst as for example described by Kishan G.,Coulier L., de Beer V. H. J., van Veen J. A. R., Niemantsverdriet J. W.,Journal of Catalysis 196, 180-189 (2000). Examples of chelating agentsare nitrilotriacetic acid, ethylenediaminetetraacetic acid (EDTA) and1,2-cyclohexanediamine-N,N,N′,N′,-tetraacetic acid.

The carrier for the catalyst is amorphous silica-alumina. The term“amorphous” indicates a lack of crystal structure, as defined by X-raydiffraction, in the carrier material, although some short range orderingmay be present. Amorphous silica-alumina suitable for use in preparingthe catalyst carrier is available commercially. Alternatively, thesilica-alumina may be prepared by precipitating an alumina and a silicahydrogel and subsequently drying and calcining the resulting material,as is well known in the art. The carrier is an amorphous silica-aluminacarrier. The amorphous silica-alumina preferably contains alumina in anamount in the range of from 5 to 75% by weight, more preferably from 10to 60% by weight as calculated on the carrier alone. A very suitableamorphous silica-alumina product for use in preparing the catalystcarrier comprises 45% by weight silica and 55% by weight alumina and iscommercially available (ex. Criterion Catalyst Company, USA).

The total BET surface area of the catalyst is preferably above 100 m²/gand more preferably between 200 and 300 m²/g. The total pore volume ispreferably above 0.4 ml/g. The upper pore volume will be determined bythe minimum surface area required.

Preferably between 5 and 40 volume percent of the total pore volume ispresent as pores having a diameter of more than 350 Å. The total porevolume is determined using the Standard Test Method for Determining PoreVolume Distribution of Catalysts by Mercury Intrusion Porosimetry, ASTMD 4284-88.

The amorphous silica-alumina carrier of the catalyst preferably has acertain minimum acidity or, said in other words, a minimum crackingactivity. Examples of suitable carriers having the required activity aredescribed in WO-A-9941337. More preferably the catalyst carrier, afterhaving been calcined, at a temperature of suitably between 400 and 1000°C., has a certain minimum n-heptane cracking activity as will bedescribed in more detail below.

The n-heptane cracking is measured by first preparing a standardcatalyst consisting of the calcined carrier and 0.4 wt % platinum.Standard catalysts are tested as 40-80 mesh particles, which are driedat 200° C. before loading in the test reactor. The reaction is carriedout in a conventional fixed-bed reactor having a length to diameterratio of 10 to 0.2. The standard catalysts are reduced prior to testingat 400° C. for 2 hrs at a hydrogen flow rate of 2.24 Nml/min and apressure of 30 bar. The actual test reaction conditions are:n-heptane/H₂ molar ratio of 0.25, total pressure 30 bar, and a gashourly space velocity of 1020 Nml/(g·h). The temperature is varied bydecreasing the temperature from 400° C. to 200° C. at 0.22° C./minute.Effluents are analysed by on-line gas chromatography. The temperature atwhich 40 wt % conversion is achieved is the n-heptane test value. Lowern-heptane test values correlate with more active catalyst.

Preferred carriers have an n-heptane cracking temperature of less than360° C., more preferably less than 350° C. and most preferably less than345° C. as measured using the above-described test. The minimumn-heptane cracking temperature is preferably more than 310° C. and morepreferably greater than 320° C.

The cracking activity of the silica-alumina carrier can be influencedby, for example, variation of the alumina distribution in the carrier,variation of the percentage of alumina in the carrier, and the type ofalumina, as is generally known to one skilled in the art. Reference inthis respect is made to the following articles which illustrate theabove: Von Bremer H., Jank M., Weber M., Wendlandt K. P., Z. anorg.allg. Chem. 505, 79-88 (1983); Léonard A. J., Ratnasamy P., Declerck F.D., Fripiat J. J., Disc. of the Faraday Soc. 1971, 98-108; and Toba M.et al, J. Mater. Chem., 1994, 4(7), 1131-1135.

The catalyst may also comprise up to 8 wt % of a large pore molecularsieve, preferably an aluminosilicate zeolite. Such zeolites are wellknown in the art, and include, for example, zeolites such as X, Y,ultrastable Y, dealuminated Y, faujasite, ZSM-12, ZSM-18, L, mordenite,beta, offretite, SSZ-24, SSZ-25, SSZ-26, SSZ-31, SSZ-33, SSZ-35 andSSZ-37, SAPO-5, SAPO-31, SAPO-36, SAPO-40, SAPO-41 and VPI-5. Large porezeolites are generally identified as those zeolites having 12-ring poreopenings. W. M. Meier and D. H. Olson, “ATLAS OF ZEOLITE STRUCTURETYPES” 3rd Edition, Butterworth-Heinemann, 1992, identify and listexamples of suitable zeolites. If a large pore molecular sieve is usedthen the well-known synthetic zeolite Y as for example described in U.S.Pat. No. 3,130,007 and ultrastable Y zeolite as for example described inU.S. Pat. No. 3,536,605 are suitable molecular sieves. Other suitablemolecular sieves are ZSM-12, zeolite beta and mordenite. Such molecularsieve containing catalysts, containing between 0.1 and 8 wt % of thesieve.

The catalyst for use in step (a) may be prepared by any of the suitablecatalyst preparation techniques known in the art. A preferred method forthe preparation of the carrier comprises mulling a mixture of theamorphous silica-alumina and a suitable liquid, extruding the mixtureand drying and calcining the resulting extrudates as for exampledescribed in EP-A-666894. The extrudates may have any suitable formknown in the art, for example cylindrical, hollow cylindrical,multilobed or twisted multilobed. A most suitable shape for the catalystparticles is cylindrical. Typically, the extrudates have a nominaldiameter of from 0.5 to 5 mm, preferably from 1 to 3 mm. Afterextrusion, the extrudates are dried.

Drying may be effected at an elevated temperature, preferably up to 800°C., more preferably up to 300° C. The period for drying is typically upto 5 hours, preferably from 30 minutes to 3 hours. Preferably, theextrudates are calcined after drying. Calcination is effected at anelevated temperature, preferably between 400 and 1000° C. Calcination ofthe extrudates is typically effected for a period of up to 5 hours,preferably from 30 minutes to 4 hours. Once the carrier has beenprepared, nickel and tungsten may be deposited onto the carriermaterial. Any of the suitable methods known in the art may be employed,for example ion exchange, competitive ion exchange and impregnation.Preferably nickel and tungsten are added by means of impregnation usinga chelating agent as described above. After impregnation, the resultingcatalyst is preferably dried and calcined at a temperature of between200 and 500° C.

Step (a) is conducted at elevated temperature and pressure. Typicaloperating temperatures for the process are in the range of from 290° C.to 430° C., preferably in the range of from 310° C. to 415° C., morepreferably in the range of from 350° C. to 415° C. Typical hydrogenpartial pressures are in the range of from 20 to 200 bar, preferably inthe range of from 70 to 160 bar. If very high viscosity index, i.e.higher than 135, base oils are desired the pressure is preferablybetween 90 to 160 bar and more preferably between 100 to 150 bar.

The hydrocarbon feed is typically treated at a weight hourly spacevelocity in the range of from 0.3 to 1.5 kg/l/h, more preferably in therange of from 0.5 to 1.2 kg/l/h. The feed may be contacted with thecatalyst in the presence of pure hydrogen.

Alternatively, it may be more convenient to use a hydrogen-containinggas, typically containing greater than 50% vol hydrogen, more preferablygreater than 60% vol hydrogen. A suitable hydrogen-containing gas is gasoriginating from a catalytic reforming plant.

Hydrogen-rich gases from other hydrotreating operations may also beused. The hydrogen-to-oil ratio is typically in the range of from 300 to5000 l/kg, preferably from 500 to 2500 l/kg, more preferably 500 to 2000l/kg, the volume of hydrogen being expressed as standard liters at 1 barand 0° C.

Step (a) and (b) may be performed in a so-called series flowconfiguration or in a two stage configuration. The series flowconfiguration is defined in that the total effluent of step (a) is usedas feed to step (b). Thus in this configuration no intermediateseparation of gaseous or liquid fractions from the effluent of step (a)takes place between steps (a) and (b). This configuration isadvantageous because it enables the process to be carried out in asimple manner, for example in one reactor or two sequentially arrangedreactors. This configuration can be advantageously applied for the cleanfeeds, such as the Fischer-Tropsch derived feed, as discussed above.Applicants have found that this configuration can also be applied formineral oil feeds suitably containing between 700 and 2000 ppm sulphurusing the above described pre-sulphided catalyst in step (a).

Another advantage of this series-flow configuration is that a gas oiland kerosene product can be separated from the effluent of step (b),which have excellent low temperature properties, such as cloud point orfreeze point.

In a series flow configuration the pressure in steps (a) and (b) will beabout the same for obvious reasons. The pressure will be determined bythe desired viscosity index as explained above. Applicants found that tomaximize the yield to base oils ex step (b) the wax conversion in step(a) is preferably between 40 and 70% and more preferably between 45 and60 wt %. The wax conversion is defined as (wax in feed to step (a)−waxin effluent of step (a))/(wax in feed to step (a))*100%.

In a two-stage configuration part of the gaseous and/or liquid fractionof the effluent of step (a) is separated off before using the effluentas feed to step (b). Preferably at least part of the ammonia andhydrogen sulphide is removed from the effluent of step (a) prior tousing said effluent as feed of step (b). Preferably the liquid effluentused as feed of step (b) comprises-after separation less than 1000 ppmwsulphur and less than 50 ppmw nitrogen. Removal of ammonia and hydrogensulphided may be achieved by well known means, for example high pressurestripping, preferably using hydrogen as stripping gas. More preferablyalso part or all of the fraction boiling substantially below the baseoil product range and the gaseous components are separated bydistillation from the effluent of step (a). An advantage of a two-stageconfiguration is that feeds having a relatively high sulphur content canbe used. A further advantage is that the pressure in step (b) can bechosen independently from the pressure in step (a). Normally one willthen consider to operate step (b) at a lower pressure. A next advantageis that because part of the compounds boiling below 370° C. can beremoved from the effluent of step (a) less feed will have to be dewaxedin step (b), enabling the use of a smaller reactor for step (b).Suitably between 5 and 40 wt % of the feed to step (a) is separated fromthe effluent of step (a) as a fraction boiling below 370° C.

The catalyst composition of the catalyst used in step (b) comprises anoble Group VIII metal, a binder and a zeolite crystallites of the MTWtype. Examples of MTW type zeolites are ZSM-12 as described in U.S. Pat.No. 3,832,449, CZH-5 as described in GB-A-2079735, Gallosilicate MTW asdescribed in Y. X. Zhi, A. Tuel, Y. Bentaarit and C. Naccache, Zeolites12, 138 (1992), Nu-13(5) as described in EP-A-59059, Theta-3 asdescribed in EP-A-162719, TPZ-12 as described in U.S. Pat. No. 4,557,919and VS-12 as described in K. M. Reddy, I. Moudrakovski and A. Sayari, J.Chem. Soc., Chem. Commun. 1994, 1491 (1994). The average crystal size ofthe zeolite is preferably smaller than 0.5 μm and more preferablysmaller than 0.1 μm as determined by the well-known X-ray diffraction(XRD) line broadening technique using the high intensity peak at about20.9 2-theta in the XRD diffraction pattern.

The binder in the catalyst may be any binder usually used for such anapplication. A possible binder includes alumina or alumina containingbinders. Applicants have found that low acidity refractory oxide bindermaterial that is essentially free of alumina provides more improvedcatalyst. Examples are low acidity refractory oxides such as silica,zirconia, titania, germanium dioxide, boria and mixtures of two or moreof these. The most preferred binder is silica. The weight ratio of themolecular sieve and the binder can be anywhere between 5:95 and 95:5.Lower zeolite content, suitable between 5 and 35 wt %, may in some casesbe advantageous for achieving an even higher selectivity.

The silica to alumina molar ratio of the zeolite prior to dealuminationis preferably larger than 50 and more preferably between 70 and 250 andmost preferably between 70 and 150. Preferably the zeolite has beensubjected to a dealumination treatment. The dealumination of the zeoliteresults in a reduction of the number of alumina moieties present in thezeolite and hence in a reduction of the mole percentage of alumina. Theexpression “alumina moiety” as used in this connection refers to anAl₂O₃-unit which is part of the framework of the aluminosilicatezeolite, i.e. which has been incorporated via covalent bindings withother oxide moieties, such as silica (SiO₂), in the framework of thezeolite. The mole percentage of alumina present in the aluminosilicatezeolite is defined as the percentage of moles Al₂O₃ relative to thetotal number of moles of oxides constituting the aluminosilicate zeolite(prior to dealumination) or modified molecular sieve (afterdealumination). Preferably dealumination is performed such that thereduction in alumina moieties in the framework is between 0.1 and 20%.

Dealumination may be performed by means of steaming. Preferably thesurface of the zeolite crystallites are selectively dealuminated. Aselective surface dealumination results in a reduction of the number ofsurface acid sites of the zeolite crystallites, whilst not affecting theinternal structure of the zeolite crystallites. When applying a surfacedealumination the reduction of alumina moieties in the framework will belower and preferably between 0.1 and 10%. Dealumination using steamresults is a typical non-selective dealumination technique.

Dealumination can be attained by methods known in the art. Particularlyuseful methods are those, wherein the dealumination selectively occurs,or anyhow is claimed to occur selectively, at the surface of thecrystallites of the molecular sieve. Examples of dealumination processesare described in WO-A-9641849. U.S. Pat. No. 5,015,361 describes amethod wherein the zeolites are contacted with sterically hindered aminecompound.

Preferably dealumination is performed by a process in which the zeoliteis contacted with an aqueous solution of a fluorosilicate salt whereinthe fluorosilicate salt is represented by the formula:(A)₂/bSiF6wherein ‘A’ is a metallic or non-metallic cation other than H+ havingthe valence ‘b’. Examples of cations ‘b’ are alkylammonium, NH₄+, Mg++,Li+, Na+, K+, Ba++, Cd++, Cu+, Ca++, Cs+, Fe++, Co++, Pb++, Mn++, Rb+,Ag+, Sr++, Tl+, and Zn++. Preferably ‘A’ is the ammonium cation. Thezeolite material may be contacted with the fluorosilicate salt at a pHof suitably between 3 and 7. Such a dealumination process is for exampledescribed in U.S. Pat. No. 5,157,191. The dealumination treatment isalso referred to as the AHS-treatment.

The catalyst composition is preferably prepared by first extruding thezeolite with the low acidity binder and subsequently subjecting theextrudate to a dealumination treatment, preferably the AHS treatment asdescribed above. It has been found that an increased mechanical strengthof the catalyst extrudate is obtained when prepared according to thissequence of steps.

It is believed that by maintaining the acidity of the catalyst at a lowlevel conversion to products boiling outside the lube boiling range isreduced. Applicants found that the catalyst should have an alpha valuebelow 50 prior to metals addition, preferably below 30, and morepreferably below 10. The alpha value is an approximate indication of thecatalytic cracking activity of the catalyst compared to a standardcatalyst. The alpha test gives the relative rate constant (rate ofnormal hexane conversion per volume of catalyst per unit time) of thetest catalyst relative to the standard catalyst, which is taken as analpha of 1 (Rate Constant=0.016 sec −1). The alpha test is described inU.S. Pat. No. 3,354,078 and in J. Catalysis, 4, 527 (1965); 6, 278(1966); and 61, 395 (1980), to which reference is made for a descriptionof the test. The experimental conditions of the test used to determinethe alpha values referred to in this specification include a constanttemperature of 538° C. and a variable flow rate as described in detailin J. Catalysis, 61, 395 (1980).

The noble metal Group VIII metal used in the catalyst of step (b) ispreferably palladium and more preferably platinum. The total amountplatinum or palladium will suitably not exceed 10% by weight calculatedas element and based on total weight of the catalyst, and preferably isin the range of from 0.1 to 5.0% by weight, more preferably from 0.2 to3.0% by weight. If both platinum and palladium are present, the weightratio of platinum to palladium may vary within wide limits, but suitablyis in the range of from 0.05 to 10, more suitably 0.1 to 5. Catalystscomprising palladium and/or platinum as the hydrogenation component arepreferred. Most preferred is when platinum is used as the solehydrogenation component. The hydrogenation component is suitably addedto the catalyst extrudate comprising the dealuminated aluminosilicatezeolite crystallites by known techniques.

The process conditions used in step (b) are preferably typical catalyticdewaxing process conditions involving operating temperatures in therange of from 200 to 500° C., preferably from 250 to 400° C., morepreferably between 300 and 380° C., hydrogen pressures in the range offrom 10 to 200 bar preferably from 30 to 150 bar, more preferably from30 to 60 bar. Weight hourly space velocities (WHSV) in the range of from0.1 to 10 kg of oil per liter of catalyst per hour (kg/l/hr), preferablyfrom 0.2 to 5 kg/l/hr, more preferably from 0.5 to 3 kg/l/hr andhydrogen to oil ratios in the range of from 100 to 2,000 liters ofhydrogen per liter of oil. The base oil having the desired viscosity canbe isolated from the effluent of step (b) by normal distillationprocesses.

The pour point of the intermediate oil product after step (a) includingany non-converted wax is suitably above 0° C. and preferably above 10°C. The pour point of the effluent of step (b) is preferably below −10°C., more preferably below −20° C. The VI is higher than 120 andpreferably higher than 130 and below 180.

It has been found that the base oil as isolated from the effluent ofstep (b) has excellent properties with respect to colour and oxidativestability. Thus an additional hydrofinishing step could be omitted.

The invention will be illustrated by the following non-limitingexamples.

Preparation of the Dewaxing Catalyst for Step (b)

MTW Type zeolite crystallites were prepared as described in “Verifiedsynthesis of zeolitic materials” as published in Micropores andmesopores materials, volume 22 (1998), pages 644-645 using tetra ethylammonium bromide as the template. The Scanning Electron Microscope (SEM)visual observed particle size showed ZSM-12 particles of between 1 and10 μm. The average crystallite size as determined by XRD line broadeningtechnique as described above was 0.05 μm. The crystallites thus obtainedwere extruded with a silica binder (10% by weight of zeolite, 90% byweight of silica binder). The extrudates were dried at 120° C. Asolution of (NH₄)₂SiF₆ (45 ml of 0.019 N solution per gram of zeolitecrystallites) was poured onto the extrudates. The mixture was thenheated at 100° C. under reflux for 17 h with gentle stirring above theextrudates. After filtration, the extrudates were washed twice withde-ionised water, dried for 2 hours at 120° C. and then calcined for 2hours at 480° C.

The thus obtained extrudate was impregnated with an aqueous solution ofplatinum tetramine hydroxide followed by drying (2 hours at 120° C.) andcalcining (2 hours at 300° C.). The catalyst was activated by reductionof the platinum under a hydrogen rate of 100 l/hr at a temperature of350° C. for 2 hours. The resulting catalyst comprised 0.35% by weight Ptsupported on the dealuminated, silica-bound MTW zeolite.

EXAMPLE 1

Example 1 illustrates a two-stage configuration as described above.

A slack wax, having a wax content of 82 wt %, a density (d70) of 0.81, anitrogen content of 30 mg/kg, a sulphur content of 766 mg/kg and aboiling range as listed in Table 1,

TABLE 1 Initial boiling point 353° C. 30 wt % 487° C. 50 wt % 510° C. 95wt % 588° C. Final boiling point 670° C.was contacted with a commercial fluorided C-454 catalyst as obtainedfrom the Criterion Catalyst Company as placed in a fixed bed reactor.The slack wax was fed to the reactor at a weight hourly space velocity(WHSV) of 1 kg/l/h. Hydrogen was fed to the reactor at an inlet pressureof 140 bar and at a flow rate of 1500 Nl/kg of feed. The reactiontemperature was 390° C.

The hydrocarbon product was distilled to remove a fraction boiling below370° C. including a gaseous fraction containing hydrogen sulphide andammonia to obtain an intermediate product having the properties aslisted in Table 2.

TABLE 2 Density 0.787 Wax content 16 wt % Kinematic viscosity at 100° C.(cSt) 5.037 Initial boiling point 366° C. 30 wt % 426° C. 50 wt % 451°C. 95 wt % 539° C. Final boiling point 611° C.

The above obtained intermediate product of Table 2 was contacted in thepresence of hydrogen with the above-described MTW zeolite containingcatalyst at an outlet pressure of 138 bar, a WHSV of 1.0 kg/l.hr and ahydrogen gas rate of 700 Nl/kg feed at various temperatures rangingbetween 320 and 345° C. The temperature was varied in order to makedifferent qualities of base oil. With different qualities of base oilsis here meant base oils having different pour points. The lowest pourpoint base oils were obtained at the highest temperature.

Gaseous components were separated from the effluent by vacuum flashingat a cutting temperature of 390° C. The oil yield and viscosity index asa function of the achieved pour point of the oil fraction (390° C.+fraction) is given in FIGS. 1 and 2. In FIG. 3 the required reactortemperature is given as a function of the resultant pour point of thebase oil.

Comparative Experiment A

Example 1 was repeated except that the intermediate product of Table 2was contacted with a catalyst based on a MFI type zeolite (ZSM-5). Thecatalyst was obtained using the same method as for the catalyst used inExample 1.

The oil yield and viscosity index as a function of the achieved pourpoint of the oil fraction (390° C.+ fraction) is given in FIGS. 1 and 2.In FIG. 3 the required reactor temperature is given as a function of theresultant pour point of the base oil.

Comparative Experiment B

Example 1 was repeated except that the intermediate product of Table 2was solvent dewaxed to base oils having pour points ranging from −37° C.to −21° C. 1 weight part of the intermediate product was diluted with 4parts of a (50/50 vol/vol) mixture of methyl ethyl ketone and toluene.This mixture was cooled to a value relating to the desired pour point ina refrigerator. The mixture was subsequently filtered at the same lowtemperature and the wax was removed from the filter. The viscosity indexof the oil and the dewaxed oil yield was measured. The oil yield andviscosity index as a function of the achieved pour point of the oilfraction (390° C.+ fraction) is given in FIGS. 1 and 2.

FIGS. 1 and 2 show that with the MTW containing catalyst base oils canbe obtained having in a higher yield as compared to when a ZSM-5dewaxing catalyst is used or solvent dewaxing is used. Furthermorehigher viscosity index values are obtained for the same base oil quality(same pour point) when the process according to the invention is used.FIG. 3 further shows that the catalyst according to the invention can beoperated at lower temperatures to achieve the same pour point reduction.Thus the MTW containing catalyst is more active. An additional advantageis that at these lower temperatures more saturation of poly-aromaticcompounds will take place. Therefore an additional hydrofinishing stepmay be omitted when the process of this invention is used.

EXAMPLE 2

Example 2 illustrates the above described series flow configurationwherein use is made of one of the more preferred catalysts for step (a):An LH-21 catalyst as obtained from Criterion Catalyst Company (Houston)was loaded into a step (a) reactor and retained as a fixed bed. TheLH-21 catalyst comprises nickel and tungsten on an acid amorphoussilica-alumina carrier and has a hydrodesulphurisation activity of 32%.The carrier of this catalyst had a heptane cracking test value ofbetween 320 and 345° C.

A slack wax, having a wax content of 65 wt %, a density (d70) of 0.804,a nitrogen content of 2 mg/kg, a sulphur content of 10 mg/kg was spikedwith dimethyldisulphide such that the total content of sulphur in thefeed was 2000 ppm. The slack wax had a boiling range as listed in Table3.

TABLE 3 Initial boiling point 347° C. 30 wt % 468° C. 50 wt % 491° C. 95wt % 561° C. Final boiling point 596° C.

The slack wax was fed to the step (a) reactor at a weight hourly spacevelocity of 1 kg/l/h. Hydrogen was fed to the reactor at an inletpressure of 145 bar and at a hydrogen flow rate of 1500 Nl/kg of feed.The reaction temperature was varied between 350 and 365° C. such toachieve a desired wax conversion as defined above.

The effluent as obtained in the step (a) reactor was as such fed to astep (b) reactor. The conditions with respect to pressure, spacevelocity and hydrogen flow rate were as in the step (a) reactor. Thereaction temperature was varied between 325 and 345° C. such to achievea desired pour point of the resultant base oils.

The base oils were obtained by cutting the effluent of the step (b)reactor at a temperature of 390° C. The yield to base oils wereexpressed as the fraction thus obtained relative to the feed to step(a). A relatively large fraction of kerosene and gas oil having goodcold flow properties were also obtained.

FIG. 4 shows the influence of the wax conversion in step (a) on theoverall base oil yield. The Figure shows that at 50 wt % wax conversionthe base oil yield is higher than at higher wax conversions in step (a).

EXAMPLE 3

The intermediate product of Table 2 was catalytically and solventdewaxed to various pour points between −20 and −37° C. using theprocedures of Example 1 and Experiment B respectively. The dynamicviscosity at −35° C. was measured according to ASTM D 2983.

The results are presented in FIG. 5. To four selected samples (see FIG.5) 0.15 wt % of an additive Plexol 154E as obtained from Rhom and Haascompany was added.

In FIG. 5 CDW relates to the catalytically dewaxed base oils, SDWrelates to the solvent dewaxed base oils and the ‘ppd’ relates to therespective base oils containing the additive. The corresponding baseoils with and without the additive are linked with an arrow.

From FIG. 5 it can be seen that both pour point and dynamic viscosity ofthe catalytic dewaxed base oils according to the present invention werereduced much more as compared to the solvent dewaxed base oilsconsisting the additive.

1. A process to prepare a lubricant having a dynamic viscosity at −35°C. of below 5000 cP by performing the following steps: (a) hydrocrackinga feed containing more than 50 wt % wax by contacting said feed in thepresence of hydrogen with a catalyst comprising a Group VIII metalcomponent supported on a refractory oxide carrier under hydrocrackingconditions sufficient to achieve between 40 and 70 wt % wax conversion;and (b) catalytically dewaxing at least part of the effluent of step (a)with a catalyst composition comprising a noble Group VIII metal, abinder and zeolite crystallites of the MTW type to obtain a base oilproduct in high yield having a pour point below −10° C. and having aviscosity index greater than 120; and (c) adding a pour point depressantadditive to the base oil product obtained in step (b) thereby obtaininga lubricant having a dynamic viscosity at 35° C. of below 5000 cP. 2.The process according to claim 1, wherein the base oil product in step(b) has a pour point below −20° C. and a viscosity index greater than130 and below
 180. 3. The process according to claim 2, wherein thenoble Group VIII metal in step (b) is platinum and the binder in step(b) is a low acidity binder which binder is essentially free of alumina.4. The process according to claim 3, wherein the binder is silica. 5.The process according to claim 4, wherein the zeolite crystallites havebeen subjected to a selective surface dealumination process.
 6. Theprocess according to claim 5, wherein the selective surfacedealumination process comprises contacting the zeolite crystallites withan aqueous solution of a fluorosilicate salt wherein the fluorosilicatesalt is represented by the formula:(A)2/bSiF6 wherein A′ is a metallic or non-metallic cation other than H+having the valence ‘b’.
 7. The process according to claim 6, wherein thewax containing feed is derived from a Fischer-Tropsch process, the GroupVIII metal in step (a) is platinum and/or palladium and wherein thetotal effluent of step (a) is used as feed to step (b) in a series flowconfiguration.
 8. The process according to claim 1, wherein the feed tostep (a) comprises at least 700 ppm sulfur, the catalyst used in step(a) is a pre-sulphided catalyst comprising a Group VIB metal and anon-noble Group VIII metal and wherein the total effluent of step (a) isused as feed to step (b) in a series flow configuration.
 9. The processaccording to claim 8, wherein the wax conversion in step (a) is between40 and 60%.
 10. The process according to claim 1, wherein the feed tostep (a) comprises between 700 and 2000 ppm sulfur, the catalyst used instep (a) is a pre-sulphided catalyst comprising a Group VIB metal and anon-noble Group VIII metal and wherein at least part of the ammonia andhydrogen sulphide which is present in the effluent of step (a) isseparated from said effluent prior to using said effluent as feed ofstep (b).
 11. The process according to claim 10, wherein the pressure instep (a) is between 100 to 150 bar and the pressure in step (b) isbetween 30 and 60 bar.
 12. The process according to claim 11, whereinthe catalyst used in step (a) is a pre-sulphided hydrodesulphurisationcatalyst comprising nickel and tungsten on an acid amorphoussilica-alumina carrier.
 13. The process according to claim 12, whereinthe sulphided hydrodesulphurisation catalyst has a hydrodesulphurisationactivity of higher than 30%, wherein the hydrodesulphurisation activityis expressed as the yield in weight percentage of C₄-hydrocarboncracking products when thiophene is contacted with the catalyst understandard hydrodesulphurisation conditions, wherein the standardconditions consist of contacting a hydrogen-thiophene mixture with 200mg of a 30-80 mesh catalyst at 1 bar and 350° C., wherein the hydrogenrate is 54 ml/mm and the thiophene concentration is 6 vol % in themixture.
 14. The process according to claim 13, wherein thehydrodesulphurisation activity of the catalyst is lower than 40%. 15.The process according to claim 14, wherein the hydrodesulphurisationcatalyst is obtained in a process wherein nickel and tungsten areimpregnated on the acid amorphous silica-alumina carrier in the presenceof a chelating agent.
 16. The process according to claim 15, wherein thealumina content of the hydrodesulphurisation catalyst is between 10 and60 wt % as calculated on the carrier alone.
 17. The process according toclaim 16, wherein the silica-alumina carrier has an n-heptane crackingtest value of between 310 and 360° C., wherein the cracking test valueis obtained by measuring the temperature at which 40 wt % of n-heptaneis converted when contacted, under standard test conditions, with acatalyst consisting of said carrier and 0.4 wt % platinum.
 18. Theprocess according to claim 17, wherein the silica-alumina carrier has ann-heptane cracking test value of between 320 and 350° C.
 19. The processaccording to claim 18, wherein the catalyst comprises between 2-10 wt %nickel and between 5-30 wt % tungsten.
 20. The process according toclaim 19, wherein the surface area of the hydrodesulphurisation catalystis between 200 and 300 m²/g.
 21. The process according to claim 20,wherein the total pore volume of the hydrodesulphurisation catalyst isabove 0.4 ml/g.
 22. The process according to claim 8, wherein the waxconversion in step (a) is between 45 and 60%.
 23. The process accordingto claim 8, wherein gas oil and kerosene product having excellent lowtemperature properties is separated from the effluent of step (b). 24.The process according to claim 1, wherein a fraction boiling below 370°C. is removed from the effluent of step (a) prior to being contactedwith the MTW type zeolite-containing catalyst composition in step (b).25. The process according to claim 24, wherein said fraction removedfrom the effluent of step (a) is between 5 and 40 wt % of the feed tostep (a).